Process for separating olefins from hydrocarbon mixtures



Feb. 24, 194s.y

115005.35 ma SEPARATING oLEFNs PROM HYDRGCARBON vNuxTurms Filed April16, 1945 Patented Feb; 24, 1948 PROCESS FOR SEPARATING OLEFIN S FROMHYDROCARBON MIXTURES Alex G. Oblad, Dallas, Tex., assigner, by mesneassignments, to Socony-Vacuum Oil Company, Incorporated, New York, N.Y., a corporation of` New York Application April 16, 1945, Serial No.588,552

1 Claim. 1

This invention relates to the separation of olens from parains. Moreparticularly this invention relates to a method for the separation oi aclose boiling fraction of low boiling hydrocarbons containing parafnsand olens into relatively pure independent streams of these classes ofhy'- drocarbons which separation cannot be accomplished successfully byconventional methods such as by fractionation o r by solvent extraction.

With the development and use of many valuable new products such asresins, synthetic rubber and Iother polymeric forms of olens, it becomeshighly desirable to produce low boiling olens in a puried state as feedmaterial to hydrocarbon synthesis processes. lA common and abundantsource of these olens is in the eilluent gases from the thermal andcatalytic cracking of petroleum oils. However, the olens are present insuch gases admixed with paraliinic gases and hence the gas must befractionated to separate cuts containing hydrocarbons of the same numberlof carbon atoms to the molecule desired in the feed to the syntheticprocess. Heretofore. these cloze cuts have been subjected to more orless conventional methods of separation such as by solvent extraction,extractive distillation or even chemical methods to separate theparalnic hydrocarbons from the oleiinic hydrocarbons. Since the olensparticularly the C3 and C4 oleflns boil very closely to thecorresponding parains, ordinary distillative fractionation methods areinadequate to produce an olefin stream of the desired purityI Solventextraction and extractive distillation only partially solve the problemsince the selectivity of known solvents for olens is not sufllcient toproduce an olenic extract yof high purity. Known chemical methods aregenerally unsatisfactory from the standpoint of yield or operating cost.

It is an object of this invention to prepare relatively pure oleflns asfeed material for such hydrocarbon synthesis processes as polymerizationand alkylation. Another object of the invention is to provide a methodfor separation of olefins from parafiinsin a narrow fraction of acracked petroleum gas stream. Still another object of the invention is'to separate a mixture of close boiling parains and olens to obtain fromsuch mixture an :olen or mixture of oleflns containingless than 5 percent of non-olenic hydrocarbons. Other obiects of the invention will beapparent from the description thereof which should be read inconjunction ywith the drawing and the appended claims.

The method of the present invention involves a simple'process ofcatalytic a'lkylation of aromatlcs by the oleflns in the parafin-olelinmixture, fractionation of the alkylate to remove therefrom non-alkylatedparains, catalytic dealkylation of the alkyl aromatics and fractionationof the dealkylated product to produce a relatively pure stream of theregenerated oleflns and a stream of aromatics for recycle tothealkylation step. Examples of hydrocarbons which can be separated by myprocess are mixtures of propylene and propane, of butanes and butenesand of pentanes and pentenes.

My process utilizes hydrocarbonsv only, namely aromatlcs such as benzeneor toluene as the re actant whereby oleilns are converted to productswhich boilat temperatures remote from the boiling temperatures of theclosely associated paraffins of the paraln-olen fraction to beseparated.

tially the same catalytic material.

In my process a stream containing paraffins and oleflns suchasbutanesand butenes is mixed with from one to twenty parts of anaromatic hydrocarbon such as benzene, and the liquid mixture is passedover an alkylatlon catalyst at alkylation temperatures to form the alkylbenzene. The

' reaction mixture is then passed to a fractionating tower forseparation of an overhead fraction of unreacted parailins from unreactedbenzene which is withdrawn as a side stream and from the f alkyl benzenewhich is removed as a bottom product. This alkyl benzene bottom productis heated to conversion temperature and passed over a 'dealkylatloncatalyst which may be the catalyst used in the alkylation step where ithas become partially spent relative to alkylation activity or it mayconsist of a different catalyst. In my preformed' ln th dealkylatlonstep. ifany, such as` toluene, are recycledto the alkylation step. Ingeneral. I prefer to operate the dealkylation step at conditions oftemperature and contact time such that no more than '75 per cent of thealkyl aromatics are dealkylated per pass through the 3 dealkylationcatalyst bed since a purer stream of the desired olefin and recycledealkylated aromatics are obtainable as a result of almost co-mpleteabsence of sde reactions when complete conversion of the alkyl aromaticsin a single pass is avoided.

I prefer to use alumina activated silica catalysts in the dealkylationstep. These may be prepared by any of the methods wellknown to thoseskilled in the art'of catalytic hydrocarbon conversion. Suitablesilica-alumina catalysts containing from,80 per cent up to 98 per centcr 99 per cent silica may be prepared by coprecipitation of silica geland alumina or by impregnating silica gel with a solution of anappropriate aluminum salt to form the alumina activated silica gel. Themixed oxides are dried and activatedy by heat treatment for severalhours at temperatures up to 200 C. prior to use.

The above catalysts are used in the dealkylation of the alkyl aromaticsin my process at temperatures within the range of from about 350 C. toabout 550 C. I prefer to use the silicaalumina catalyst at temperatureswithin the range of from about 475 C. to about 550 C. when operating atspace velocities within the range of from about to 20 volumes of alkylaromatics (liquid basis) per volume of catalyst space per hour. Ifdesired, somewhat lower temperatures may be used, i. e., from about 400C. to 500 C., at space velocities of from 0.2 to 5 volumes of alkylaromatics (liquid basis) per volume of catalyst space per hour. l

The above catalysts may also be used in the alkylation step oi! myprocess at lower temperatures and generally higher space velocities.Thus, the silica-alumina type catalyst may be used to alkylate thearomatics with the olefins at temperatures within the range of fromabout 175 C. to 300 C., at space velocities within the range of fromabout 1 to 10 volumes of mixed olens and aromatics (liquid basis) perhour while temperatures up to 350 C. or 400 C. may be used if spacevelocities from up to 50 volumes of feed per volume of catalyst spaceper hour are used in vapor phase operation. When operating my process tostore exothermic heat of alkylation in the catalyst bed, as describedbelow, after the catalyst has lost appreciable alkylation activity thetemperature of operation of the alkylation step is -higher than themaxima of the above ranges. Thus, a maximum of 400 C. for normaloperation in the alkylation step would be increased to 425 C. or even450 C. in order to store heat in the catalyst and thereby lower the heatduty of furnace 36 in preheating alkylate prior to the dealkylationstep.

Catalysts other than the above may be used for the alkylation step of myprocess. Thus, Friedel-Crafts type catalysts such as anhydrous hydrogenfluoride, hydrogen halide promoted aluminum halide or H3PO4 impregnatedalumina and H3PO4 impregnated kleselguhr may be used in the alkylationstep. I prefer to use silicaalumina catalyst in the alkylation stepsince this catalyst may also be used in the dealkylation step afterbeing partially spent with respect to alkylation activity or vice versamay be used in the alkylation step after use in the dealkylation step.However, I do not wish to limit my invention to the use of any speccalkylation catalyst or 'any speciilc dealkylation catalyst, and I do notwish to be limited to the use of the same catalyst in these two steps ofmy process.

The alkylation step of my process may be carrled out in either liquidphaseor vapor phase at pressures within the range of from 100 to 1500pounds gage although I prefer to operate in vapor phase using thealumina-silica catalyst at relatively high temperature and high spacevelocity. The dealkylation step is carried out in the vapor phase atpressures from atmospheric up to fty pounds gage.

Referring now to the drawing, which is a diagrammatic illustration ofone method of carrying out the process of my invention, a mixture ofbenzene and a closely fractionated C4 cut from cracked gasolinecontaining n-butane, isobutane, butene-2, butene-l and isobutane, theratio of benzene to C4 hydrocarbons in said mixture being at least l to1, is passed by means of pump I0 through line II, heat exchanger I2' andline I3 to heat supply means I 4 where the mixture is heated to atemperature of about 250 C. The heated mixture is then passed via lineI5 to manifold line I6 and thence through valve I1 in line I8 toalkylation reactor 20 at a pressure of about 300 pounds per square inch.

Reactors 20, 40, and 60 are stationary bed reactors of the heatexchanger type the tubes of which are packed with alumina-silicacatalyst of the hereinabove described coprecipitate type. Thesereactors, which may alternatively consist of towers containing catalystin a continuous bed or towers containing catalyst disposed on a seriesof trays, are aternately put on the alkylation cycle, on thedealkylation cycle, and on the catalyst regeneration cycle of theprocess. As the mixed reactants pass upward through the catalyst intower 20 at least a part of the benzene is alkylated with' the oleiinsof the cracked fraction and the product stream iiows through line 2I andheat exchanger I2 and the liquid product is passed by means of pump 23through line 24 to alkylate fractionator 25. The conversion oi' benzeneto alkylated benzene is an exothermic reaction, and hence a coolingmedium may be circulated through reactor 20 on the shell side in orderto remove the exothermic heat of the reaction. This cooling medium maybe water or any other suitable cooling fluid or it may consist of thebottom products from fractionation towers 25 and I6 described below. Iidesired, the cooling medium may be circulated through the tube sectionof the reactor. the catalyst being packed in the shell section.

Fractionator 25 which is equipped with reflux means 26 and reboilermeans 21 is operated at a pressure of about 100 pounds per square inchand serves to separate the non-alkylated parafnic components of the C4hydrocarbons from benzene and alkylated benzene. The parallnic C4fraction containing not more than 5.0 per cent and preferably less than0.3 per cent by weight of non-alkylated C4 olens is passed overhead fromfractionator 25 through line 28. This whole product may be used as feedto a butane isomerization process or the butane fraction may be so usedafter isobutane is removed therefrom by fractionation in cases whereinthe isobutane content is relatively high. If desired, the residualoleins, if any, may be removed by a sulfuric acid wash before passingthe paramnic C4 fraction to isomerization. The use of excess benzene inthe alkylatlon step insures a substantially complete olefin cleanup fromthe gas stream, and, hence, additional puriilcation of the paraflinicstream in line 28 is usually unnecessary. Unreacted benzene is Withdrawnfrom fractionator 25 as a side stream for recycle to the alkylation stepthrough line 2'9 which leads to line Il, and the alkyl benzene bottomproduct is withdrawn through line 30.

While reactor 20 is being utilized for the alkylation cycle, reactors 40and 60 areon stream for the dealkylation and catalyst regenerationcycles, respectively. Alkylated benzene in line 30 is passed to line 32and thence to the shell side of reactor 60 where the product picks upheat sup plied by the regeneration of catalyst in the tube sectionofreactor 60 thereby aiding in the control of the temperature ofregeneration of the catalyst. If desired, a part of the alkyl benzeneproduct may be diverted through valved line 33 to the shell side ofalkylation reactor 20 in order to remove the exothermic heat ofalkylation. This preheated stream passes from reactor 20 through lines34 and 35 to furnace 36 for nal heating prior to conversion in reactor40. Preheated alkyl benzene passes from the shell section of regenerator60 via line 31 which connects with line 34 leading to furnace 36.

In furnace 36 the alkylated product is brought to a temperature withinthe range of from 350 C. to 550 C. preferably from 475 C. to 525 C. Thisproduct is then passed as a vapor through lines 38 and 39 to thecatalyst packed tubes of reactor 40 which has been previously on streamfor the alkylation cycle. Reactant vapors pass upward through the tubesof reactor 40 at a space velocity (liquid basis) within the range offrom about to 20 volumes of reactant per 6 ample, I may operate myprocess by using reactor 20, the catalyst in which has been freshlyregenerated, for the alkylation step. Reactor 40 is then on thedealkylation cyclehaving been shifted thereto from the alkylation cyclewithout an intermediate regeneration, and the catalyst bed in reactor60, freshly regnerated on a relatively short cycle and cooled asdescribed hereinbelow, is standing by ready for use on the alkylationcycle. After the catalyst in reactor is partially spent the alkylationfeed stream in line I6 is diverted through line 52 toreactor 60,preheated alkylate in line 38 is seht to reactor 20, and the catalyst inreactor 40 is put on the regeneration cycle. When changing the flow ofreactants to different towers suitable adjustment of valves in lines 32,33, and 54 are made to redirect the flow of alkylate from fractionatorand recycle from fractionator 46 to As stated hereinabove, I prefer tooperate the dealkylation cycle to obtain less than 75% conversion, and,hence, the dealkylate product will contain olens, predominantly butenes,benzene and non-converted alkyl benzenes. Fractionator 46, which isoperated at a pressure of about 100 pounds per square inch, is equippedwith suitable reflux means 4l vand reboiler means 48. The olefinfraction is taken overhead through line 49 and recycle benzene is eitherrecovered as a side stream through line 62 or is recycled through line50 which connects with process feed line Il. The bottom product fromfractionator 46 which contains unconverted alkyl benzene with a smallamount of olefin polymer ls withdrawn through line 6| which connectswith line 32 and the stream after receiving preheat in catalystregcnerator 60 and nal heat pick up in the furnace 36 is recycled toreactor 40.

During at least a part of the period when reactors 20 and 40 `are beingutilized for the alkylation and dealkylation cycles, the catalyst inreactor 60, which reactor has been removed from operation on a previousdealkylation cycle, and, therefore, contains carbon-deactivatedcatalyst, is being subjected to oxidative regeneration. The tendency forcarbon to collect 0n the catalyst in either the alkylation ordealkylation cycles is not pronounced, and, hence, the sum of the cycleperiods for these two operations will be relatively long compared to thetime required for regeneration of the catalyst. Thus, for exthe shellside of the appropriate reactors for the pickup of preheat from theexothermi'c alkylation and regeneration cycles.

The inequality of time requirement for the conversion cycles and theregeneration cycle permits a unique sequence of operations which aids inthe reduction of heat duty for furnace 36. Thus, as an alternate methodof supplying heat to the endothermic dealkylation reaction zone, I mayoperate my process in the following manner. When the catalyst in reactor20 which is on stream for alkylation begins to show loss of activity, Ipermit the temperature of the catalyst bed to rise as much as '75 or 100C. above the alkylation temperature ranges described hereinabove. 'I'hisis accomplished by reducing or eliminating entirely the withdrawal ofthe exothermic heat of reaction by coolants on the shell side of thereactor. Product from reactor `20 is then directed through cooler 56 inline 5l which connects with manifold alkylation feed line I6 whence theproduct is directed to freshly regenerated catalyst in reactor 60 wherethe alkylation is completed at a lower temperature and lower pressure asdescribed hereinbelow.

The two stage alkylated product is passed from reactor 60 through line59 which connects with manifold alkylate product line 6| which in turnjoins line 2 I. If desired, the ratio of benzene to butene in thealkylation feed may be increased during this period in order tocompensate for reduced alkylation due to lower catalystactivity andunfavorably high temperatures in reactor 20. Excess benzene may bewithdrawn from the sys# tem through trapout line 62, leading from frac-.tionator 46. Thus, in the succeeding cycle when the catalyst in reactor40 is more completely spent the dealkylation feed in line 3B is directedto the heated catalyst in reactor 20, alkylation feed is sent directlyto reactor 60 and reactor 40 is put on the regeneration cycle.

When operating according to the above procedure, that is, at abnormallyhigh temperature in the alkylation zone, the reaction may be morecompletely directed toward alkylation by adjustr ment of the relativeamount of aromatic hydrocarbons to olefin hydrocarbons in the alkylationfeed and/or by adjustment of the pressure in the alkylation reactor.Thus. for example, the ratio of aromatics to olelns in the alkylationfeed may be raised when operating in the higher temperature range from 3or 5 to 1 to a ratio'as high as 9 or 10 to 1. A second method,collaterally or alternately used for increasing the alkylation underhigh temperature operation consists of allowing the pressure to build upin reactor 20 from a, normal operating pressure of `200 or 300 poundsper square inch to a. pressure of '700 or 800 pounds or even 1500 poundsper square inch. Thus, although the temperature inthis type of operationis above the most efficient level for conventional moderate pressure andlower aromatic to olefin ratio alkylation procedure, alkylation willpredominate rather than dealkylation because of the above describedadjustment of pressure and the relatively high aromatic to olen ratio.

The catalyst in reactor 40 is regenerated by burning off the depositedcarbon by means of a gas containing free oxygen such as a mixture of airand iiue gas. It is necessary to control the oxidation in order toprevent the temperature of the catalyst undergoing reactivation fromrising above 650 C. and preferably the temperature of the catalystshould not be allowed to exceed about 600' C. during the catalystregeneration period. Hence, air, diluted with from 2 to 20 parts of aninert gas such as flue gas is introduced to the catalyst in reactor 40through manifold line 63 and line S4. The oxygen in the gaseous mixturecombines with the deposited carbon to form carbon dioxide and carbonmonoxide, and the catalyst is thereby regenerated to its originalactivity. The regenerator gas passes from reactor 40 through line 65which connects with manifold line 66 through which the gas passes fromthe system. If desired; at least a part of the eiliuent gas from'reactor40 may be recycled as a diluent gas for the oxygen containingregenerator feed gas by passing a part of the gas Vfrom line 66 throughvalved line 61 and cooler 68. Cooler 68 may be substituted by anindirect heat exchanger such as heater I4, thereby utilizing a part ofthe sensible heat of the flue gas as preheat for alkylation chargestock.

Following the regeneration of the catalyst in reactor 40 the flow ofcold ue gas is continued for sumcient time to lower the temperature ofthe catalyst to the approximate required alkyiation temperatures,assuming that the freshly regenerated catalyst is to be used immediatelyfor the alkylation cycle. If the reactor is to be used in the succeedingcycle for dealkylating the product in line 5I, cooling of the bed willbe unnecessary. In any case, the catalyst in reactor 40 should be purgedwith an inert gas such as flue gas to remove all traces of free oxygenbefore inauguration of either of the hydrocarbon conversion cycles. Thereactor should also be purged with flue gas prior to the regenerationcycle in order to remove hydrocarbon vapors from the reactor beforeintroduction of free oxygen containing gas in the regeneration cycle.Alternately, steam may be used to purge the spent catalyst beforeregeneration and/or after regeneration before being subjected to theabove hydrocarbon conversion cycles.

Although I have described one embodiment of the invention whereinstationary -beds of catalyst are used in the' three cycles of operation,I do not wish to be limited to stationary bed type operation. Theprocessmay be carried out very readily using moving beds of catalystsuch as are used in the Thermofor catalytic conversion processes. thecatalyst being moved continuously from the alkylation zone to thedealkylation zone and thence to regeneration whence it would be recycledto the alkylation zone. The process can also be carried out in hinderedow type operation utilizing powdered catalyst or combinations of thesediierent methods of operating catalytic C. Isobutane, boiling point 11.7Isobutylene, boiling point 6.9 Butene-l, boiling point 6.3 n-Butane,boiling point .6 Eutelia-2, (trans) boiling point .86 Butene-2, (cis)boiling point 3.64

The separation of these C4 hydrocarbons in a reasonable degree of purityby distillative fractionation is impossible. However, by the hereinabovedescribed process the normal butane and isobutane may be recovered as arelatively pure paraflinic mixture as overhead from fractionator and thenormal butane may then be separated from the isobutane by simpledistillative fractionation. The olens of the mixture are recovered asoverhead from fractionator 46, and this mixture is subjected tofractionation to obtain an isobutylene-butene-l cut and a butene-Z, cut,the latter serving as ideal feed to a dehydrogenation process for theproduction of butadiene. The isobutylene can be separated from thebutene-l by-absorption in 65% H2804 from which it may be recovered as apure isobutylene feed to a BFS catalyzed polymerization process or othercatalytic process for the production of high molecular weight polymers.The butene-l may be used in the production of intermediates for themanufacture of various chemicals.

My process is likewise adaptable for the isolation of relatively pureolefin streams from a C: or Cs fraction of petroleum cracked gases.

While I have shown and described the preferred embodiment of myinvention, I Wish it to be understood that I do not confine myself tothe precise details herein set forth by way of illustration, as it isapparent that many changes may be made therein by those skilled in theart, without departing from the spirit of the invention, or exceedingthe scope of the appended claim.

I claim:

A process for the separation of oleflns of from 3 to 5 carbon atoms fromclose-boiling parafns and for separately recovering said olensand saidparaiilns which comprises the steps of t (1) admixing a molar excess ofbenzene on the basis of the olefin charged with the paraffin-olefinmixture and subjecting th'e feed mixture so obtained to a pre-alkylationstep by contacting with an alumina-silica catalyst in a first zone at atemperature above 300 C. and at a pressure of from about '700 to about1500 pounds per square inch while permitting the temperature within saidfirst zone to rise due to the exothermic heat of reaction, (2) coolingthe gaseous eiiluent from step 1 below 300 C. and subjecting it to afinal alkylation step by contacting with an alumina-silica catalyst in asecond zone while maintaining a temperature of from 175 C. to 300 C. anda pressure of less than about 300 pounds per square inch, (3)fractionating the eiiiuent from step 2 to obtain a paraffin stream, andunreacted benzene stream and an alkyl benzene stream (4l subjecting thealkyl benzene to a dealkylation step by contacting it with analumina-silica catalyst in a third zone maintained at a temperatureabove 350 C.. (5) fractionating the eilluent from step 4 to obtain anolen stream, a benzene stream and an unreacted alkyl benzene stream, (6)recycling the unreacted alkyl benzene stream from step 5 to step 4andthe benzene streams from steps 3 and 5 to step 1, (7) separatelyrecovering the paraffin stream and the olen stream, (8) when thetemperature in the rst zone has risen to about 450 C., transferring thedealkylation step to the rst zone, passing the feed mixture directly tothe nai alkylation step in the second zone and regenerating the catalystin the third zone to prepare it for reuse as an alkylation catalyst, and(9) when the catalyst has been regenerated, placing it in use incarrying out the nal 15 alkylation step and raising the pressure of thereactants in the second zone to make the operations therein thepre-alkylation step and permitting the temperature to rise to repeat thecycle. ALEX G. OBLAD.

REFERENCES CITED The following references are of record in the file .ofthis patent:

10 UNITED STATES PATENTS Number Name Date 1,953,702 Davidson ---s Apr.3, 1934 2,222,632 Sachanen et al Nov. 26, 1940 2,242,960 Sachanen et alMay 20, 1941 2,295,608 Ruthruff Sept. 15, 1942 2,360,358 Mattox Oct. 17,1944 2,370,810 Morrell et al. Mar. 6, 1945 l 2,381,175 Mattox Aug. 7,1945 0 2,382,505 Schulze Aug.14,1945

FOREIGN PATENTS Number Country Date 456,637 Great Britain Nov. 12, 1936OTHER REFERENCES Babor and Lehrman, General College Chem- 20 Thomas etal., "Hydrocarbon Reactions-IV.

Removal of Side Chains from Aromatics, J. A. C. S. 66, 1694-5-6, October1944.

Ethyl Benzene-Outlet, Oil and Gas Journal, August 6, 1942, pages 14 and15.

